Multicatalyst dehydrogenation, single catalyst decontamination and aromatization



Oct. 10, 1961 P. EVANS 3,003,948 MULTICATALYST DEHYDROGENATION, SINGLECATALYST DECONTAMINATION AND AROMATIZATION Filed Dec. 19, 1958 TORECOVERY 0 39 v 93 1a 34 56 i0 IUI 55 R'EFORMATEI INVENTOR mwmm't R wm ANT United States Patent 3,003,948 MULTICATALYST DEHYDROGENATION, SINGLECATALYST DECONTAMINATION AND AROMA- TIZATION 1 Louis P. Evans, Woodbury,N. J., assignor to Socony Mobil Oil Company, Inc., a corporation of NewYork Filed Dec. 19, 1958, Ser. No. 781,789 6 Claims. (Cl. 208-65) Thepresent invention relates to the decontamination of naphthas and thereforming of the decontaminated naphthas over a platinum-group metalreforming catalyst and, more particularly to the decontamination over afirst catalyst, dehydrogenation over a platinumgrou-p metal catalyst,and aromatization of the dehydrogenated naphtha over the aforesaid firstcatalyst. More specifically, the present invention relates todecontamination of a naphtha in the presence of a first particle-formsolid catalyst having hydrogenating and desulfurizing capabilities asWell as aromatizing capabilities and preferably, in addition, thecapability of dehydrogenating naphthenes and aromatizing paraflins,dehydrogenating the decontaminated naphtha over a static bed ofplatinumgroup metal reforming catalyst, and completing the reforming ofthe so-treated naphtha to produce the required octane over the aforesaidfirst catalyst. Preferably, the aforesaid first catalyst is employed inaccordance with moving bed or fluidized technique while theplatinum-group metal reforming catalyst is employed as a static bed.

It has been known for some time that in the reforming of naphtha over aplatinum-group metal catalyst in a multi-bed system at pressures of 500p.s.i.g. or more that most of the carbon is deposited in those reactorsin which reactions other than the dehydrogenation of naphthenes takeplace. Thus, in a three reactor system wherein dehydrogenation ofnaphthenes take place primarily in the first reactor, the coke isdeposited primarily in the third reactor. That is to say, the cokedeposited upon the catalyst in the third reactor amounts to about 25 percent based on the weight of catalyst, whereas the amount of cokedeposited on the same amount of catalyst in the first reactor, where theprincipal reaction is dehydrogenation of naphthenes, is less than fivepercent. From this it follows that, when the reactions other thandehydrogenation of the naphthenes are carried out in accordance with,moving bed or fluidized technique the on stream time can be increased anindustrially important amount.

Furthermore, it, is known that when reforming naphtha at pressures lessthan 500 p.s.i.g., e.g., at 15 to 250 p.s.i.g. the deposition of coke isgreater than when reforming to the same octane rating under pressures ofthe order of 500 p.s.i.g. or more. But even at pressures less than 500p.s.i.g. there is more coke deposited in those beds of catalyst in whichthe principal reactions are others than dehydrogenation of naphthenes.object of the present invention to provide a method of reforming naphthaat pressures less than 500 p.s.i.g., preferably at pressures of about 15to about 250 p.s.i.g., in which dehydrogenation occurs principally inthe presence of one or more static beds of platinum-group metalparticle-fonn reforming catalyst while the reactions of dehydrocyclizationor aromatization and hydrocracking occur principally in the presence ofa moving bed or fluid: ized bed of base metal oxide reforming catalyst.Other objects and advantages of the present inventionwill be,- comeapparent to those skilled in the art from the following description ofthe present invention taken in conjunctionwith the illustrative butnon-limiting flowsheet of the drawing which illustrates one embodimentofv the.

present invention. When the efiiuent from the pretreating section of thefirst reactor contains not more than 1 p.p.m. (part per million) ofnitrogen calculated on the naphtha there is no need to remove theammonia. However, when the efiluent from the pretreatin-g section of thefirst reactor containsmore than 1 p.p.m. of nitrogen suf-' ficientammonia must be removed therefrom to reduce the nitrogen content of thefeed to the second and third reactors to not more than 1 p.p.m. ofnitrogen calculated on the naphtha.

The present invention will be described hereinafter first for apretreating section eifiuent which does not contain more than 1 p.p.m.of nitrogen calculated on the naphtha and then for a pretreating sectionefiluent which does contain more than 1 p.p.m. of nitrogen calculated onthe naphtha.

Thus, a naphtha containing an amount of nitrogen which can be reduced tonot more than 1 p.p.m. in the presence of the catalyst employed andunder the reaction conditions in the pretreating section of reactor 23,for example a naphtha charge stock containing not more than about 10p.p.m. of nitrogen when using a mixture of oxides of cobalt andmolybdenum is contacted with a moving bed of catalyst havinghydrogenating and hydrodesulfiurizing capabilities together with thecapae bilities of dehydrocyclizing and hydrocracking paraffinsaIllustrative of this type of catalyst is a catalyst comprising oxides ofone or more metals selected from the group consisting of chromium,molybdenum, tungsten, cobalt, nickel, silicon, zirconium, titanium,magnesium and cerium at least one of which reversibly combines withsulfur. The presently preferred catalysts which can be used for thispurpose are (A) a mixture of oxides of chromium, aluminum andmolybdenum, and (B) a mixture of oxides of cobalt and molybdenum. Theuse in the first stage of a catalyst at least one component of whichreversibly combines with the sulfur liberated from the naphtha feed hasthe advantage that the generated sulfur does not contaminate thehydrogen-containing gas and hence reuse of the hydrogen-containing gasdoes not require a separate gas recycle purifying circuit. Consequently,the hydrogen-containing recycle gas can be circulated through not onlythe reactor of the first stage but also through the reactor(s) of thesecond stage.

The hydrodesulfurization and reforming of the hydrodesulfurized naphthais carried out at pressures of about 15 to about 250 p.s.i.g.,preferably about 100 to about 200 p.s.i.g.

In the flow sheet of the invention an embodiment of the inventionhas-been selected for the purpose of illustration in which feed naphthais contacted in the presence of hydrogen with a moving bed ofparticle-form solid re- Accordingly, it is an forming catalyticmaterial. "(A fluidized bed of catalyst can'be used in place of themoving bed of catalyst in the illustrative flow sheet.) I Y Thetemperature to which the feed naphtha is heated in coil 13 in heater 14is dependent upon (1) the required reaction temperature in section D(desulfurization section) of reactor 23, (2) the temperature to whichthehydrogen-containing gas is heated in coil 20 in heater 21, and (3) themol ratio of hydrogen-containing gas tofeed naphtha. Whenhydrodesulfurization is the principal re-- action in section B ofreactor 23 the reaction temperature therein is between about 600 andabout 850' F. It is desirable to heat the feed naphtha to thelowesttemperature at which the required temperature for-decontaminzktion can bemaintained in section D at the designed catalyst-to-oiljratios While heating the hydrogen-containing gas to the'. maximumpractical temperature. Thus, the

feed naphtha; is heated to a temperaturewithin the-range,

of about 55.05 to about 800 F, and the molar ratioofthehydrogen-containing gas to feed naphtha is about 0.1 to about 3.0. Underthese conditions the hydrogen-containing gas is heated to a temperaturewithin the range of about 1000 to about 1200 F.

However, section D can also be operated under conditions such that aportion of the naphthenes in the feed naphtha is dehydrogenated and thefeed naphtha is decontaminated in section D of reactor 23. For thispurpose the temperature in section D is within the range of about 800 toabout 1000 F. The temperature of the feed naphtha entering section D iswithin the range of about 770 to about 970 F. The mol ratio of thehydrogencontaining gas to feed naphtha is about 0.5 to about 5.0; andthe hydrogen-containing gas is heated to a temperature within the rangeof about l000 to about 1200 F.

Accordingly, feed naphtha drawn from a source not shown through pipe 1by pump 2 is discharged into pipe 3. The feed naphtha flows through pipe3 and branches 5 and 6 under control of valves 7 and 8 respectively toabsorber 4. The quantities of feed naphtha flowing through branches 5and 6 is controlled as hereinafter described. In absorber 4 the feednaphtha is contacted with a stripping gas to remove water, oxygen, anddeposit precursors. For this purpose, a hydrogen-containing gas isemployed. Since the feed naphtha will absorb gasoline hydrocarbons froma gas it is preferred to contact the feed naphtha withhydrogen-containing gas flowing from liquid-gas separator 16 throughconduit 18 and under control of valves 104 and 107. Accordingly,hydrogen-containing gas flows from liquid-gas separator 16 throughconduits 18, 105, and 106 into absorber 4. The amounts of feed naphthaflowing through pipes 5 and 6 and the amount of gas flowing fromseparator 16 through conduits 105 and 106 is balanced to removesubstantially all of the water, oxygen and deposit precursors from thefeed naphtha and to remove substantially all of the gasoline, i.e., Cand heavier hydrocarbons from the gas.

The stripped gas flows from absorber 4 through conduit 9 to the refineryfuel gas main or to hydrogenation processes. The enriched feed naphthaflows from absorber 4 through pipe 10 to heat exchanger 11.

In heat exchanger 11 the enriched feed naphtha is in indirect heatexchange relation with the effluent from the third or aromatizingsection A of reactor 23 flowing through conduit 96. From heat exchanger11 the enriched feed naphtha flows through pipe 12 to coil 13 in heater14. In coil 13 in heater 14 the enriched naphtha is heated to adecontaminating or naphthene dehydrogenating reaction temperature withinthe range of about 550 to about 970 F. dependent upon the requiredtemperature in section D of reactor 23, the temperature to which thehydrogen-containing gas is heated in coil 20 in heater 21, and therecycle gas to naphtha mol ratio.

The heated enriched feed naphtha flows from coil 13 in heater 14 throughconduit 15 to vapor inlet 64 of section D of reactor 23.Hydrogen-containing gas heated to a temperature within the range ofabout 900 to about 1200 F. in coil in heater 21 flows through conduit 22under control of valve 91 to a point in conduit 15 intermediate heater14 and reactor inlet 64. In conduit 15 the heated hydrogen-containinggas is mixed with the aforesaid heated enriched feed naphtha in theratio of about 500 to about 1500 s.c.f. of hydrogen per barrel of feednaphtha or in the ratio of about 0.6 to about 1.8 mols of hydrogen permol of feed naphtha. The charge mixture so formed flows through conduit15 to inlet 64 of the decontaminating section D of reactor 23.

Reactor 23 is divided into two reaction zones, i.e., first reaction zoneD and third reaction zone A and two sealing zones. Decontaminating zoneD is formed by plates 24 and 25. A sealing zone is formed immediatelybeneath decontaminating zone D between plates 25 and 26. Aromatizingzone A is formed immediately below the aforesaid'disengaging zonebetween plates 26 and- 27.

The second sealing zone is formed immediately below zone A betweenplates 27 and 28.

The flow of catalyst through reactor 23 will be described first andthereafter the flow of vapors through reactor 23 will be described.

A catalyst comprising up to about 18 percent chromium oxide and thebalance at least 72 percent alumina impregnated with about 3 to about 15percent molybdenum is employed in reactor 23.

When reactor 23 is operated under a pressure in excess of about p.s.i.g.the catalyst transfer means illustrated in the drawing is employed. Foroperations at pressures between about 15 and about 100 p.s.i.g. a feedleg such as described in US. Patent No. 2,829,087, issued April 1, 1958,can be employed.

Accordingly, since it will be assumed that reactor 23 is operated at apressure of about p.s.i.g., a catalyst transfer means such asillustrated is employed. Hot active catalyst in feed hopper 44 flowsunder control of valve 45 into pressuring chamber 46 (valve 47 closed).Pressuring chamber 46 is purged (valves 45 and 47 closed) with any inertgas such as flue gas, nitrogen or the like by means of an inlet and avent not shown. After purging, the pressure in chamber 46 is raised tothat of decontaminating section D or preferably 5 to 10 pounds higherwith hydrogen-containing gas, preferably recycle gas, from gas-liquidseparator 16. When the pressure in chamber 46 has been raised to atleast the pressure in decontaminating zone D, the pressuring gas isshut-0ff and valve 47 opened. The catalyst flows from pressuring chamber46 into reactor 23 onto distributing plate 26. Valve 47 is closed andpressuring chamber purged with an inert gas. After purging pressuringchamber 46 valve 45 is opened, chamber 46 filled with catalyst and thecycle repeated.

Plate 26 is provided with a plurality of orifices each of which issubtended by a catalyst conduit 29. The catalyst flows from plate 26downwardly through conduits 29 into decontaminating section D. Indecontaminating section D the catalyst column is annular in shape andconfined between foraminous partition 30 and tubular foraminouspartition 31. This provides a central inlet conduit 32 and a peripheralcollecting zone 33.

The zone immediately below decontaminating zone D is adapted to permitcatalyst to flow from decontaminating zone D into aromatization zone Awith no substantial loss of vapors from zone D into zone A where zone Ais at a pressure up to 5 p.s.i. less than the pressure in zone D. Thesealing means comprises a plurality of conduits represented by conduits34. Alternatively, conduits 34 do not extend to plate 26 and recycle gasis introduced into the space between plate 25 and plate 26 to provide asubstantially total seal.

The hot catalyst flows through conduits 34 into aromatization zone A. Inzone A' the catalyst flows as an annular column confined betweenforaminous partition 35 and tubular foraminous partition 36. Foraminouspartition 35 being tubular provides means 37 for introducing reactantsinto the catalyst bed. An annular collecting zone 38 is formed betweenforaminous partition 36 and the liner or inside of the shell of reactor23.

The catalyst flows downwardly as a substantially continuous annularcolumn of particle-form solid catalytic material through zone A to thecatalyst draw-off zone immediately below aromatization zone A. Thecatalyst is drawn off by any suitable means such as conduits 39 as morefully described in US. Patent No. 2,412,135. The catalyst flowsdownwardly through conduits 39 into collector 40 and thence undercontrol of regulator 49 into surge chamber 50 (valve 51 closed).Depressuring chamber 52 is purged with an inert gas such as fiue gas andthen pressured with hydrogen-containing gas such as recycle gas by meansnot shown. After pressuring chamber 52 to about the pressure ofaromatization zone A valve components of the zone D efiluent.

51 is opened (valve 53 closed) and the catalyst in surge chamber 50flows into depressuring chamber 52. Valve 51 is closed and chamber 52depressured to atmospheric pressure (by means not shown). Chamber 52 isthen purged with an inert gas. Valve 53 opens and the catalyst flowsfrom chamber 52 into surge chamber 54. From surge chamber 54 thecatalyst flows down chute 55 to means feeding a catalyst transfer means56 such as a bucket elevator, a pneumatic lift or the like whereby thecatalyst is elevated to chute 57. The catalyst flows down chute 57 tokiln hopper 58 and thence into kiln 59.

Kiln 59 is of any suitable type wherein the coke can be burned off thecatalyst in a stream of combustionsupporting gas such as air. Kilns ofthis type are wellknown to those skilled in the art and have beendescribed in several publications including US. Patent Nos. 2,726,994and 2,724,683. From kiln 59 the regenerated catalyst flows throughconduit 60 to chute 61. The regenerated catalyst flows down chute 61 tomeans feeding a catalyst transfer device 62 such as a bucket elevator, apneumatic lift or the like. Elevator 62 discharges the regeneratedcatalyst into chute 63 down which the regenerated catalyst flows toreactor feed hopper 44.

Returning now to conduit 15; the charge mixture of enriched feed naphthaand hydrogen in the proportions set forth hereinbefore and at atemperature within the range of about 600 to about 1000 F. dependentupon the considerations set forth hereinbefore flows through conduit 15to decontaminating zone vapor inlet 64. The charge mixture flows fromvapor inlet 64 to conduit 32 formed by the tubular foraminous partition31. From conduit 32 the vapors of the charge mixture flow radiallythrough the foraminous partition 31 into the annular substantiallycontinuous substantially compact column of first or decontaminatingcatalyst and through the foraminous partition 30 into annular collectingzone 33. From collecting zone 33 the vapors now designated zone Defiiuent flow through zone D outlet 65 to conduit 66 having valve 67 andbranch 68 with valve 69.

When the zone D effluent, as was set forth hereinbefore, contains notmore than 1 ppm. of nitrogen it is usually unnecessary to separate thenon-condensible components of the zone D efiluent from the naphtha orcondensible Accordingly, the zone D eflluent flows through conduit 66(valve 69 closed; valve 67 open) to reactor manifold 70.

Since the reaction temperature and space velocity in zone D were thoseunder which little or no dehydrogenation of naphthenes occurred in zoneD, in accordance with the principles of the present invention reactionconditions in reactors 75 and 76 are such as to dehydrogenate at least85 percent of the naphthenes present in the effiuent from zone D. (Theeffluent from zone D hereinafter is designated pretreated feed naphtha.)

While reactors 75 and 76 are illustrated as radial flow adiabaticreactors those skilled in the art will understand that adiabaticreactors designed for vertical flow as Well as isothermal reactors canbe used.

In the illustrated radial flow reactors the bed of catalyst is confinedin the form of an annulus byperipheral foraminous partitions 81 and 82and central tubular foraminous partitions 77 and 78 in reactors 75 and76 respectively. An annular collecting zone is provided in each reactorbetween the peripheral foraminous partition 81 or 82 and the liner orshell of the reactors 75 or 76 respectively. The annular collecting zonein reactor 75 is designated 83 and that in reactor 76 is numbered 84.

The catalyst in reactors 75 and 76 is a platinum-type catalyst, e.g.,comprising about 0.35 to about 2 percent by weight of a metal of theplatinum group on a refractory oxide carrier such as alumina,silica-alumina, silica or the like with or without about 0.1 to about0.8 percent of a halogen, .i.e., chlorine, fluorine, bromine.

The reaction conditions for fresh or fresh regenerated catalyst are asfollows:

1 Limited by the maximum practical reaction temperature to which thecatalyst can be exposed without irreversible deactivation and usuallyabout 20 F. less than the aforesaid maximum practical reactiontemperature.

' In the drawing two static bed reactors piped for parallel flow of thepretreated naphtha feed are shown. While one reactor can be used, it ispreferred to use two for easeof construction, operation and maintenance.I

The pretreated naphtha feed, the unused hydrogen and the products ofdecontamination other than sulfur (it is preferred to use in the reactor23 a catalyst such as a molybdenum oxide-chromium oxide-alumina catalystwhich forms a metal sulfide with the sulfur produced by hydrogenation ofthe sulfur-containing components of the naphtha feed) and containing notmore than about 1 ppm. of nitrogen based on the naphtha feed, flowthrough conduit 66 to manifold 70. At some point on conduit 66intermediate valve 67 and manifold70 hydrogen-containing gas, preferablyhydrogen-containing recycle gas flowing from heater 21 through conduit22 to conduit and thence to conduit 66, is mixed with the efliuent ofzone D containing not more than 1 ppm. of nitrogen based upon thenaphtha feed in an amount to provide, with the hydrogen in the zone Deffluent, a hydrogen-to-naphtha mol ratio of about 1 to 10. Thereforming charge mixture thus provided flows through conduit 66 tomanifold 70 and manifold branches 71 and 72 under control of valves 73and 74 respectively to reactors 75 and 76.

The reforming charge mixture flows into central conduits 79 and 80,radially through the tubular foraminous partions 77 and 78 into thecatalyst beds in reactors 75 and 76. The reaction products flow from thecatalyst beds through the foraminous partitions 81 and 82 into theannular collection zones 83 and 84 of reactors 75 and 76 respectively.The reaction products flow from collecting zone 83 (reactor 75) throughreactor outlet 85 into conduit 87 and thence to conduit 89. The reactionproducts flow from collecting zone 84 (reactor 76) to reactor outlet 86to conduit 88 and thence to conduit 891 The components of the reactionproducts in conduit 89 boiling in the gasoline range comprise at leastabout 85 percent of the naphthenes of the naphtha feed as aromatichydrocarbons together with substantially all of the paraflins of thenaphtha feed as paraflins.

The effluents of the static bed reactors, now desig: nateddehydrocyclization feed, have a temperature of about 800 to 950 F. Whennecessary additional hydrogen-containing gas, preferably recycle gasflowing from heater 21 through conduits 22, 9 0 and 93 under control ofvalve 94 is mixed with the dehydrocyclization feed in conduit 89 inamount required to provide a dehydrocyclization charge mixturetemperature of about 900l050 F. The dehydrocyclization charge mixtureflows through conduit 89 to reactor inlet 95 of aromatization zone A ofreactor 23.

From aromatization zone A inlet 95 the dehydrocyclization charge mixtureflows upwardly through central conduit 37 and radially therefrom throughforaminous partition 35 and the catalyst bed and foraminous parti tion36 to annular collecting zone 38. From annular collecting zone 38 thereaction products of the aromatization and hydrocracking of thedehydrocyclization feed flow through aromatiz-ationzone A vapor outlet76 to con duit 96. The aromatization zone efliuent, now desig natedfinal eflluent, flows through conduit 96 to heat exchanger 11 where thefinal eflluent is in indirect heat exchange relation with thenaphthafeed flowing from absorber 4 through pipe 10 as describedhereinbefore.

From heat exchanger 11 the final etfiuent flows through conduit 97 toheat exchanger 98 where the final efiluent is in indirect heat exchangerelation with the condensate flowing from liquid-gas separator 16through pipe 99. From heat exchanger 98 the final efliuent flows throughconduit 100 to cooler 101 where the final eflluent is cooled to atemperature at which under the existing pressure C and heavierhydrocarbons are condensed. The cooled final efiluent flows from cooler101 through conduit 102 to liquid-gas separator 16. In liquid-gasseparator 16 the uncondensed components of the final efiluent areseparated from the C and heavier hydrocarbons.

The uncondensed final efiluent, now designated recycle gas andcomprising hydrogen, traces of ammonia and C to C hydrocarbons flowsthrough conduit 17 to conduit 18. A portion about equal to the gas makein reactors 23, 75 and 76 is diverted through conduit 103 to conduits105 and 106 as previously described herein. The balance of the recyclegas flows through conduits 18 and 19 to coil 20 in heater 21 asdescribed hereinbefore.

The condensed portion of the final efiluent, now designated condensate,comprising 0., and heavier hydrocarbons and traces of hydrogen andammonia, flows from separator 16 through pipe 99 to heat exchanger 98 aspreviously described herein. From heat exchanger 98 the condensate flowsthrough pipe 108 to fractionator 109.

In fractionator 109 C and lighter hydrocarbons and other components ofthe condensate volatile at temperatures not higher than the boilingpoint of C hydrocarbons are taken overhead through pipe 110 to means forthe recovery of C hydrocarbons. A portion of the bottoms in fractionator109, comprising hydrocarbons heavier than C flows through a re-boilercomprising pipe 111, heat exchanger 112 and pipe 113. Any means forheating the bottoms of fractionator 109 to a temperature at which Chydrocarbons are volatile can be used in place of the re-boilerillustrated. The net product portion of the bottoms of fractionator 109flows through pipe 114 for finishing, e.g., addition of tetraethyl lead,anti-iclng additive, etc. as high octane gasoline. 1

When the naphtha feed contains more than 1 ppm. of nitrogen and anitrogen-sensitive catalyst is employed in any of the reactors 75 and76, provision must be made to remove sufficient of the ammonia producedin zone D of reactor 23 from the effluent of zone D to ensure that thecharge mixture entering reactors 75 and 76 does not contain more than 1ppm. of nitrogen based on the naphtha feed.

The molybdenum oxide-chromium oxide-aluminum oxide catalyst preferredfor decontamination of the feed naphtha and aromatization of thedehydrocyclization feed is substantially insensitive. to nitrogen.Accordingly, the charge naphtha is heated, mixed withhydrogen-containing gas and introduced into decontamination zone D ofreactor 23 as previously described herein. However, since the chargenaphtha contains more than 1 p.p.m. of nitrogen, the efiluent of zone Dcontains more than 1 ppm. of nitrogen and perforce the excess nitrogenmust be removed from the zone D etfiuent before contacting thepretreated charge naphtha with the nitrogen-sensitive platinum-typecatalyst.

Accordingly, with valve 67 closed and valve 69 open the zone D effiuentflows through conduit 68 to heat exchanger 116 where the zone D eflluentis in indirect heat exchange relation with the bottoms of splitter 130flowing from pump 146 through pipe 147. From heat exchanger 116 the zoneD eflluent flows through conduit 117 to heat exchanger 118 where thezone D effluent is in indirect heat exchange relation with the zone Dcondensate flowing from liquid-gas separator 122 through pipe 128. Fromheat exchanger 118 the zone D efiluent fiows through conduit 119 tocooler 120 where the zone D effiuent is cooled to a temperature at whichC and heavier hydrocarbons are condensed at the existing pressure. Fromcooler the zone D effluent flows through conduit 121 to liquid-gasseparator 122. In liquid-gas separator 122 the uncondensed zone Detlluent is separated from the condensed zone D efl'luent. Theuncondensed zone D effiuent, designated impure recycle gas, flows fromseparator 122 through conduit 123 to conduit 136.

The impure recycle gas flows through conduit 136 to means 126 forremoving ammonia, e.g., dilute sulfuric acid. From means for removingammonia 126 the purified recycle gas flows through conduit 127 toconduit 19.

The condensed portion of the zone D eflluent, now designated zone Dcondensate, flows from separator 122 through pipe 128 to heat exchanger118 and thence through pipe 129 to fractionator 130. In fractionator 130a temperature is maintained at which C to C hydrocarbons, ammonia andother components of the zone D condensate are volatile. The ammonia andlight hydrocarbons are taken as an overhead through conduit 131 tocooler 132 where the overhead is cooled to a temperature at which C andheavier hydrocarbons are condensed. The overhead flows from cooler 132through conduit 133 to accumulator 134. In accumulator 134 the C andlighter hydrocarbons, hydrogen, and ammonia separate from the C andheavier hydrocarbons of the overhead and flow through conduit 135 toconduit 136 where the accumulator overhead mixes with the impure recyclegas. In the event that there is a net make of gas in zone D the net gasmake can be vented through conduit 124 under control of valve 125 toabsorber 4. The condensed overhead flows from accumulator 134 throughpipe 137 to the suction side of pump 138. Pump 138 discharges thecondensed overhead into pipe 139. A portion of the condensed overheadflows through pipe 139 to fractionator 130 for use as reflux. Thebalance flows through pipe 140 under control of valve 141 for recoveryas light naphtha.

A portion of the bottoms of fractionator 130 fiows through the re-boilersystem comprising pipe 142, heat exchanger 143 and pipe 144 or any othersuitable means for heating the fractionator bottoms to the temperaturenecessary to maintain in fractionator 130 a temperature at which C andlighter hydrocarbons are volatile.

The balance of the bottoms of fractionator 130 flows through pipe 145 tothe suction side of pump 146. Pump 146 discharges the balance of thebottoms into pipe 147 through which the fractionator bottoms, nowdesignated pretreated naphtha flows to heat exchanger 116. In heatexchanger 116 the pretreated naphtha is in indirect heat exchangerelation with the zone D eflluent as described hereinbefore. From heatexchanger 116 the pretreated naphtha flows through pipe 148 undercontrol of valve 149 to coil 150 in heater 151. In coil 150 thepretreated naphtha is heated to a reaction temperature within the rangeof about 800 to about 1000 F. dependent upon the amount of recycle gasand the temperature of the recycle gas flowing through conduit 90 to bemixed with the pretreated naphtha in conduit 66. From coil 150 thepretreated naphtha flows through pipe 152 to conduit 66. The pretreatednaphtha and the hydrogen-containing gas, preferably recycle gas flowingthrough conduit 90, mixed therewith in the ratio of about 1 to 10 molsof hydrogen per mol of naphtha to form a charge mixture flows throughconduit 66 to manifold 70 and thence to reactors 75 and 76 ashereinbefore described.

Those skilled in the art will understand, from the description of theillustrative examples of the present invention given hereinbefore thatthe present invention provides a method of pretreating and reformingnaphtha wherein the naphtha is contacted with a first particleform solidreforming catalyst having hydrodecontaminating capabilities in thepresence of hydrogen in a first reaction zone, the eflluent of saidfirst zone containing not more than 1 ppm. of nitrogen based upon saidnaphtha when contacted with the platinum-group metal reforming catalyst,is contacted in at least one second reaction Zone th a econd reformingcatalyst comprising a platinum-group particle-form solid reformingcatalyst to convert a preponderant portion, preferably at least 85percent of the naphthenes present in said naphtha to aromatics, and theeffluent of said second reaction zone is contacted in a third reactionzone with the aforesaid first reforming zone. That the reactionconditions in the first reaction zone are decontaminating ordecontaminating and dehydrogenating conditions. That the reactionconditions in the second reaction zone are dehydrogenating conditions.That the reaction conditions in the third zone are those necessary fordehydrocyclization, or dehydrocyclization and hydrocracking, ofparafiins. That the efliuent of the third reaction zone is separatedinto a liquid portion having the required octane and a gaseous portioncomprising hydrogen and light hydrocarbons which is recycled to saidfirst, second and third reaction zones. It is preferred that thereforming catalyst having hydrodecontaminating capabilities used in thefirst and third reaction zones react with the sulfur in the feed naphthain the presence of hydrogen to form a sulfide of one of the componentsof the catalyst stable under the conditions of temperature and hydrogenpartial pressure existing in the hydrodecontaminating Zone, and that thesulfide thus formed be oxidizable to a gaseous sulfur compound under theconditions under which the aforesaid catalyst is regenerated. Thoseskilled in the art will understand that it is preferred to pass thefirst reforming catalyst through the first and third reaction zones as asubstantially compact annular column of catalyst, and that the amount offirst reforming catalyst and the temperature thereof is proportioned tothe amount of charge mixture, i.e., charge naphtha plushydrogen-containing gas, and the temperature thereof to maintain therequired hydrodecontaminating, or hydrodecontaminating and naphthenedehydrogenating, temperature in the first reaction zone and the requireddehydrocyclizing, or dehydrocyclizing and hydrocracking, temperature inthe third reaction zone. The first reforming catalyst enters the firstreaction zone at a temperature of about 600 to about 1000 F., preferablyabout 750 to about 850 F. The first catalyst enters the third reactionzone at a temperature Within the range of about 850 to about 1200 F.,preferably about 900 to about 1000 F.

I claim:

1. A method of pretreating and reforming naphtha which comprises flowinghot, active, nitrogen-insensitive, sulfur-insensitive particle-formfirst reforming catalyst successively through a first reaction zone anda third reaction zone, said particle-forrn first reforming catalysthaving hydrodecontaminating capabilities, at least one component of saidfirst reforming catalyst forming a sulfide stable under the conditionsof temperature and hydrogen partial pressure in said first reaction zoneand oxidizable to a gaseous sulfur compound during regeneration, in saidfirst reaction zone contacting said flowing first reforming catalystwith feed naphtha containing more than 1 p.p.m. of nitrogen at atemperature of at least 800 F. to obtain a first reaction zone effluentcomprising ammonia, hydrogen sulfide, hydrogen, and C and heavierhydrocarbons, in a second reaction zone contacting at least the C andheavier hydrocarbons of said first reaction zone efliuent withparticle-form nitrogen-sensitive platinum-group metal in the presence ofhydrogen and not more than 1 p.p.m. of nitrogen based upon the naphthafeed to said first reaction zone under reforming conditions oftemperature, pressure, and liquid hourly space velocity to dehydrogenateat least 85 percent of the naphthenes in the naphtha feed to said firstreaction zone to aromatics to obtain a second reaction zone efiiuentcomprising hydrogen, and C and heavier aromatics 10 a and paraffins, inthe aforesaid third reaction zone, without separation of any substantialportion of said reaction zone efiiuent, contacting substantially all ofsaid sec- 7 0nd reaction zone effluent with said particle-form firstreforming catalyst discharged from said first reaction zone atsubstantially paraffin dehydrocyclizing temperature under paraffindehydrocyclizing conditions of temperature, pressure, andcatalyst-to-oil ratio to dehydrocyclicize paraflins in said secondreaction zone effluent and to substantially complete dehydrogenation ofnaphthenes in said second reaction zone effluent to obtain a thirdreaction zone efltluent, cooling said third reaction zone efiluent to atemperature at which C and heavier hydrocarbons condense, separatinguncondensed third reaction zone efiiuent comprising hydrogen and C to Chydrocarbons from condensed C and heavier hydrocarbons, and recyclingsaid uncondensed third reaction zone efiiuent to at least said firstreaction zone.

2. The method of pretreating and reforming naphtha as set forth anddescribed in claim 1 wherein the temperature in said first reaction zoneis a naphthene dehydrogenating temperature.

3. The method as set forth and described in claim 1 wherein the firstreforming catalyst comprises oxides of chromium, molybdenum, andaluminum.

4. The method of pretreating and reforming naphtha as set forth anddescribed in claim 1 wherein the naphtha feed to the first reaction zonecontains more than 10 p.p.m. of nitrogen, wherein the first reactionzone efiiuent' is separated into (1) ammonia-containing recycle gas. (2)C and heavier hydrocarbons, wherein said ammoniacontaining recycle gasis treated to remove ammonia, wherein said treated recycle gas isrecycled to said first reaction zone, wherein said C and heavierhydrocarbons are separated into light naphtha and second reaction zonefeed, and wherein said second reaction zone feed is mixed with at leasta portion of the heated uncondensed third reaction zone efiiuent.

5. The method of pretreating and reforming naphtha as set forth anddescribed in claim 1 wherein the hot, active, nitrogen-insensitive,sulfur-insensitive particle-form first reforming catalyst flows as asubstantially compact annular column downwardly through said first andthird reaction zones, wherein the amount and temperature of said firstreforming catalyst and the amount and temperature of charge mixturecomprising feed naphtha and hydrogen are proportioned in said firstreaction zone to maintain at least hydrodecontaminating temperature insaid first reaction zone, and wherein the amount and temperature of saidfirst reforming catalyst and the amount and temperature of the secondreforming zone efliuent are proportioned in said third reaction zone tomaintain at least paraffin dehydrocyclicizing temperature in said thirdreaction zone.

6. The method of pretreating and reforming naphtha as set forth anddescribed in claim 1 wherein the first reforming catalyst comprisesoxides of chromium, molybdenum, and aluminum, wherein the first reactionzone effluent contains not more than 1 p.p.m. of nitrogen based on thefeed naphtha introduced into the first reaction zone, and whereinsubstantially the whole of the first reaction zone efiiuent flowsdirectly to the second reaction zone.

References Cited in the file of this patent UNITED STATES PATENTS2,573,149 Kassel Oct. 30, 1951 2,629,683 Haensel Feb. 24, 1953 2,664,386Haensel Dec. 29, 1953 2,758,062 Arundale et a1. Aug. 7, 1956 2,758,064Haensel Aug. 7, 1956

1. A METHOD OF PRETREATING AND REFORMING NAPHTHA WHICH COMPRISES FLOWINGHOT, ACTIVE, NITROGEN-INSENTIVE, SULFUR-INSENSITIVE PARTICLE-FORM FIRSTREFORMING CATALYST SUCCESSIVELY THROUGH A FIRST REACTION ZONE AND ATHIRD REACTION ZONE, SAID PARTICLE-FORM FIRST REFORMING CATALYST HAVINGHYDRODECONTAMINATING CAPABILITIES, AT LEAST ONE COMPONENT OF SAID FIRSTREFORMING CATALYST FORMING A SULFIDE STABLE UNDER THE CONDITIONS OFTEMPERATURE AND HYDROGEN PARTIAL PRESSURE IN SAID FIRST REACTION ZONEAND OXIDIZABLE TO A GASEOUS SULFUR COMPOUND DURING REGENERATION, IN SAIDFIRST REACTION ZONE CONTACTING SAID FLOWING FIRST REFORMING CATALYSTWITH FEED NAPHTHA CONTAINING MORE THAN 1 P.P.M. OF NITROGEN AT ATEMPERATURE OF AT LEAST 800*F. TO OBTAIN A FIRST REACTION ZONE EFFLUENTCOMPRISING AMMONIA, HYDROGEN SULFIDE, HYDROGEN, AND C1 AND HEAVIERHYDROCARBONS, IN A SECOND REACTION ZONE CONTACTING AT LEAST THE C5 ANDHEAVIER HYDROCARBONS OF SAID FIRST REACTION ZONE EFFLUENT WITHPARTICLE-FORM NITROGEN-SENSITIVE PLATINUM-GROUP METAL IN THE PRESENCE OFHYDROGEN AND NOT MORE THAN 1 P.P.M. OF NITROGEN BASED UPON THE NAPHTHAFEED TO SAID FIRST REACTION ZONE UNDER REFORMING CONDITIONS OFTEMPERATURE, PRESSURE, AND LIQUID HOURLY SPACE VELOCITY TO DEHYDROGENATEAT LEAST 85 PERCENT OF THE NAPHTHENES IN THE NAPHTHA FEED TO SAID FIRSTREACTION ZONE TO AROMATICS TO OBTAIN A SECOND REACTION ZONE EFFLUENTCOMPRISING HYDROGEN, AND C5 AND HEAVIER AROMATICS AND PARAFFINS, IN THEAFORESAID THIRD REACTION ZONE, WITHOUT SEPARATION OF ANY SUBSTANTIALPORTION OF SAID REACTION ZONE EFFLUENT, CONTACTING SUBSTANTIALLY ALL OFSAID SECOND REACTION ZONE EFFLUENT WITH SAID PARTICLE-FORM FIRSTREFORMING CATALYST DISCHARGED FROM SAID FIRST REACTION ZONE ATSUBSTANTIALLY PARAFFIN DEHYDROCYCLIZING TEMPERATURE UNDER PARAFFINDEHYDROCYCLIZING CONDITIONS OF TEMPERATURE, PRESSURE, ANDCATALYST-TO-OIL RATIO TO DEHYDROCYCLICIZE PARAFFINS IN SAID SECONDREACTION ZONE EFFLUENT AND TO SUBSTANTIALLY COMPLETE DEHYDROGENATION OFNAPHTHENES IN SAID SECOND REACTION ZONE EFFLUENT TO OBTAIN A THIRDREACTION ZONE EFFLUENT, COOLING SAID THIRD REACTION ZONE EFFLUENT TO ATEMPERATURE AT WHICH C4 AND HEAVIER HYDROCARBONS CONDENSE, SEPARATINGUNCONDENSED THIRD REACTION ZONE EFFLUENT COMPRISING HYDROGEN AND C1 TOC3 HYDROCARBONS FROM CONDENSED C4 AND HEAVIER HYDROCARBONS, ANDRECYCLING SAID UNCONDENSED THIRD REACTION ZONE EFFLUENT TO AT LEAST SAIDFIRST REACTION ZONE.